Hydrocarbon conversion



\ Aug. 7, 1945. E, L.. D'ouylLLE ETAL 2,381,439

HYDROCARBON CONVERSION Filed DeC. 15, 1941 www. W

WWRNRNNTWN II-velz Patented Aug. 7, 1945 IIYDROCABBON CONVERSION Edmond L. douvnle and Bernard L. Ewing,

Chicago, lll., assignors to Standard Oil Company, Chicago, lll., a corporation of Indiana Application December 15, 1941, Serial No. 422,984-

4 Claims. (Cl. 260-683.5)

This invention relates to an improved process vfor the isomerization of low-boiling parainic hydrocarbons by means of an aluminum halide catalyst activated by a hydrogen halide.' This application is a continuation-impart of our copending application No. 245,570 which issued as Patent No, 2,266,012 on December 16, 1941. l

An object of our invention is to provide an improved system for the eflcient and more complete conversion of parafllnic hydrocarbons to isomers of higher octane number. Another object of our invention is to provide an improved method for the introduction of catalyst into a conversion zone. A further object of our invention is to provide improved methods for effecting intimate contact between catalyst, light paraiiinic hydrocarbons, activator and hydrogen in a reaction zone. A still further object of our invention is to provide a method for obtaining the greatest possible yield of isoparamnic hyrocarbons per unit of catalyst consumed in an isomerization process. Other objects and advantages will become apparent as the detailed description of our invention proceeds.

When our invention is applied to refinery naphtha. we find it desirable to first fractionate the naphthas in order to remove undesirable components; for the isomerization of a. charging I liquid straight-chain or slightly branched-chain paraiiinhydrocarbon, such as normal lpentane, normal hexane, 2 methyl pentane, etc., but generally, predominantly `straight-run naphthas such as those from Michigan, Pennsylvania and lVIid-Continent crude oils are preferable, since they are much more readily available. Another excellent feed stock is a highly parailinic naphtha produced by the Fischer-Tropsch process from carbon monoxide and hydrogen. Natural gasoline fractions and so-called distillates are also suitable and are plentiful and inexpensive in some production areas. 'I'he feed stocks are preferably in the butane to hexane boiling range and they mayhave end points (or 95% points) as high as 170 to 180 F. although lower 95% points are preferred. They preferably contain less than about and preferably 0.5' to 2.0% or less of aromatic hydrocarbons. In many cases a preliminary solvent extraction step or other treatment is necessary or vdesirable in order to reduce the aromatic content.` Olefinic hydrocarbons are also undesirable and should not be present in more than very small amounts, while cycloparafflns or naphthenic hydrocarbons can be tolerated in considerable quantities. It is also frequently desirable to limit or eliminate parajmnic hydrocarbons higher boiling than the hexanes. The heptane content should not exceed 10% and should preferably be less than 5%, The feed stocks should preferably contain at least 50% of parain hydrocarbons and those .containing at least of paraffin hydrocarbons are especially desirable. At least a substantial part of the charging stock is employed for absorbing hydrogen chloride from gases produced in the system.

The aluminum chloride catalyst is preferably introduced as a solution in a hydrocarbon stock, and most desirably in a butane fraction. It is possible however to employ a light naphtha cut containing a substantial amount of butane as the solvent, or even to use thelCs--C fraction for this purpose. The part of the charge containing the hydrogen chloride is passed through a heat exchanger after the absorption step and then introduced into a reactor. v'I'he part of the charge used for dissolving make-up aluminum' chloride is heated prior to the solution step in order that the desired amount of solution-may be effected. Hydrogen should be added to the reactor in the case of Cs-Cs charging stocks. The reactor is operated at a temperature Within the approximate range of from F. to 400 F., preferably 2.00" F. to 300 F., and at a pressure Within the approximate range of about 100 to about 3000 pounds .per square inch, preferably about 500 to about 1500 pounds per square inch. For a Cs-Cs fraction we have found that excellent results are obtainable at a temperature within the approximate vicinity of 250 F. and a pressure within the approximate vicinity of about 850 to 900 pounds per square inch.

'I'he amount of hydrogen required will vary somewhat with temperature, pressure and hydrogen chloride concentration in the reaction zone, ranging from about 20 cubic feet per barrel of stock charged at low temperatures, low pressures and low hydrogen chloride concentrations to 200 or more cubic feet per barrel at high temperatures, pressures and high hydrogen chloride concentrations. For optimum conditions in the operation on a pentane-hexane hydrocarbon fraction, the actual vhydrogenveonsumption will be about 100 cubic feet per barrel and in order to insure.the presence of the requisite amount of hydrogen in the reactor we prefer to introduce from about 100 to about 300, preferably about 200 cubic feet of hydrogen per barrel of charging stock.

The amount of make-up aluminum chloride may be within the approximate range ofy 0.1 to 4.0 pounds per barrel of total charging stock, usually within the general vicinity of 1 pound per barrel. An important feature of our invention is the introduction of this make-up aluminum chloride to the reaction zone in the form of a solution in a light hydrocarbon, which hydrocarbon may be a portion of the original feed stock, but more desirably is a fraction rich in butane. l

,The amount of hydrogen chloride may be within the approximate range of about 5 to about 30 pounds of hydrogen chloride per barrel of total stock charged, usually within the general vicinity of about pounds per barrel, but since only a small portion of the hydrogen chloride is actually consumed, the major portion of it may be recovered and reintroduced as will be hereinafter described.

While the materials introduced into the reactor with the charging stock are aluminum chloride,

-hydrogen chloride, and hydrogen, the effective catalyst in the reactor is an aluminum chloridehydrocarbon complex which may contain more or less dissolved or uncombined aluminum chloride. At the beginning of an operation we prefer to have the reaction towers at least about half full of this complex although the complex can be formed in situ by the combination of hydrocarbons with aluminum chloride in the ;resence of hydrogen chloride.V .The fresh complex is relatively non-viscous and has a specific gravity about twice as great as that of the charging stock, so that when charging stock is introduced at the base of the reactor it flows as a dispersed phase upwardly through the complex, thus effecting intimate contact between the charging stock and the complex. The incoming charging stock is mainly in the liquid phase but it may be partially vaporized by the gaseoushydrogen, about half of the volume of the upflowing stream being gaseous because of the introduced hydrogen which serves the function of promoting turbulence and effecting intimate mixing of charging stock and complex. Based on the stock charged and the total amount of complex in the reactor, the space velocity should be within the approximate range of 0.2 to 4 volumes of liquid feed per hour per volume of complex in the reactor, preferably about 1/2 to2 volumes per `hour per volume of complex'.

In the upper part of the reactor sumcient settling space is provided to permit separation of complex from the clear reaction product. To insure. adequate settling weA can either increase the cross-sectional area of the upper part of the reactor or provide a separate chamber of large, horizontal, cross-sectional area. This enlarged upper section or separate chamber has two important functions: (1) It provides a soaking drum wherein dissolved catalyst can effect further isomerization of the substantially clear product for obtaining a more desirable product distribution. and (2) it provides a large crosssectional area for substantially .complete separation of any entrained complex so that the complex can be returned to the reactor ybefore the products enter the cooler. The settled complex is returned to the reactor or to v`a separate recovery system and the clear products are passed through a cooler to a low-pressure settling chamber for the release of hydrogen chloride and dissolved catalyst.

At room temperature and atmospheric pressure only about 0.10 gram of aluminum chloride can be dissolved in a liter of light naphtha (or isomerization product) of the type described. Under the reaction temperature and pressure about 50 to 75 times as much aluminum chloride may be held in solution. It is essential that dissolved catalyst material, whether aluminum chloride or complex, be removed from the products before they enter the fractionation system, not only because of catalyst losses which would otherwise be suffered but because of the complications and expense that would be encountered in handling a product containing such dissolved catalyst material. The gases containing hydrogen chloride can be scrubbed by a portion of the charging stock for recovering the hydrogen chloride and the substantially catalystfree product can be charged to a hydrogen chloride stripper and sent to a fractionation system.

We have found that a considerable amount of hydrogen chloride can be recovered from the product in a simple stripping. column which is operated at a pressure of about 200'pounds per square inch, a top temperature in the approximate range of 100 F. to 150 F. and a bottom temperature in the approximate range of 300 F. to 400 F. The final products are neutralized with caustic, water washed and, if necessary or desirable, they can be fractionated or stabilized.

We can employ a multi-stage reaction with complex, the first stage being with relatively spent complex at relatively high temperature, and the second step with relatively fresh complex at a lower temperature. Complex from the low temperature stage can be transferred to the high temperature stage. Substantially constant complex activity Ycan be maintained in each zone by the addition of active materials thereto and the withdrawal of relatively spent catalyst therefrom. Hydrogen chloride can be recovered from the spent catalyst by treatment with Water or sulfuric acid.

' The invention will be more fully understood from the following detailed description of a spe- I I3 which can be about 41/2 feet in diameter by about 50 to 55 feet high. This fractionator is provided with conventional heating means I4 at its base, and is operated under such conditions that heptane and heavier hydrocarbons are withdrawn from the base of the column through line I5. Where a close fractionation is desired between butanes and lheavier hydrocarbons or between hexanes and lighter hydrocarbons a plurality of fractionating columns will of course be employed but in the accompanying drawing such a fractionation is diagrammatically illustrated by a side stream draw-off on fractionating column I3. Such a side stream draw-017 may of course be actually used `where close fractem through line 2|. and valved line 22, while the lighter gases are vented overhead. The pentanes and hexanes can be withdrawn as a side stream by pump 23 through cooler 24 and line 25,'and about 50% to about 100% of this stream introduced through line 26 to the top of absorber 21. Any charging stock not sent to the absorber can be directed to the solution tank by pump 28 and line 29.

The hydrogen chloride required for the reaction is absorbed in the maior portion of the feed stock, preferably the' higher boiling portion thereof, before it is heated or introduced into the reactor. The hydrogen chloride absorber can be about one and one-half feet in diameter by twenty-eight feet high and it can operate at av pressure of from about 100 to about 300, for example about 250 pounds per square inch. A stream of hydrogen chloride gases from the system is introduced at the base of this absorber through line 3| Make-up hydrogen chloride can up hydrogen chloride we can employ chlorine, an alkyl chloride, or other substance which will supply the necessary hydrogen halide activator under reaction conditions. We prefer, however, to employ hydrogen chloride and to generate it, if necessary, in a separate generator.

The'hydrogen chloride generator 33 can be of any known type. The chlorine supplying agent introduced through line 34 is preferably chlorine gas although it may be sodium chloride, 22 muriatic (hydrochloric) acid or other halogen-containing agent. The hydrogen-supplying agent introduced through line 35 can be hydrogen gas, a hydrocarbon, sulfuric acid, etc. 'I'hus hydrogen and chlorine can be burned in generator 33 to supply hydrogen chloride. Wax tailings or other hydrocarbons can be introducedthrough line 35 and chlorinated by chlorine gas introducedby line 34 to produce hydrogen chloride and chlorinated hydrocarbons.v (Additional hydrogen chloride can of course be obtained from the latter.) Sodium chloride or hydrochloric acid can be introducedv through line 34 and sulfuric acid through line 35, but in our system this hydrogen lchloride generator can operate under such pressure that. no compressors are required for introducing the hydrogen chloride throughline 36 to the base of absorber '21, and this hydrogen chloride does not require the purification which is generally necessary even for the productionA of commercial grades of hydrochloric acid. By-

small amounts of hydrogen, methane, etc.,are purged from the system through line 38, thus eliminating not only gaseous impurities from line 3| but also gaseous impurities from line 36.

It is a. special feature of this invention that the fresh aluminum chloride be introduced into the system in solution in a light hydrocarbon stock. At room temperature and pressure, the solubility of aluminum chloride in ordinary naphtha is negligible and only abut 0.1 gram of aluminum chloride can be dissolved ln a liter of the Cav-Ce hydrocarbons. Under the reaction temperatures and pressures, however, about 50 to 75 times as much aluminum chloride can be held in solution be introduced through line 32. Instead of make- 4 products from the hydrogen chloride generator Y are withdrawn through line 31.. Y

The hydrogen chloride'picked up in absorber l 21 should be suilicient to give an amount of hydrogen chloride in the total stock entering the reactors within the approximate range ofy 2% to 10%; that is, in the general vicinity of 5% by weight based on total stock charged. Two-thirds or morev of this hydrogen chloride can be obtained 'buygases from line 3|. Unabsorbed gases such as in the Cfr-Ce hydrocarbons. Aluminum chloride has been found to be much more soluble in butanes, and particularly in normal butane. so that in a preferred embodiment of our invention we prefer to dissolve the aluminum chloride in a butane stream and to add the butane with its dissolved catalyst to the light naphtha stream going to the reactors in order to supply make-up aluminum chloride. For example, at about 250 F. approximately 5.5% by weight of aluminum chloride will be dissolved in normal butane. Moreover, the presence of normal butane in the isomerization system is not at all detrimental and in the upper range o f temperatures employed therein is beneficial, due to the fact that it represses the formation of any additional amounts of normally gaseous hydrocarbons of this type. In addition, at least a part of the normal butane Will undergo isomerization to isobutane-a product valuable in itself for the synthesis of various high octane number fuels and for blending with gasolines of low volatility in limited amounts, Accordingly, butane or a fraction rich in butane, from reflux drum I8 may be pumped by pump 39 through line 40 and heater 4| to aluminum chloride solution tank 42 at a temperature of about 250 F. or thereabouts. This hot stream dissolves sucient aluminum chloride to supply all of the necessary make-up aluminum chloride to the systern.- Fresh aluminum chloride is added from source 43 to aluminum chloride solution tank 42 and the aluminum chloride may be present in the tank in the form of lump or powdered anhydrous aluminum chloride and may `be distributedA throughout the tower or maintained in beds or by other similar means. Suilicient pressure will of course be required to maintain the -butane fraction in liquid state and the pressure may be high enough to introduce the solution into the reaction zone. The solution tank is diagrammatically illustrated in -the drawing and in actual practice may consist of one or more upow chambers wherein aluminum chloride lumps are retained on a suitable screen support, the butane stream kentering the tank at the bottom'and leaving at the top. One of such solution tanks may be charged with aluminum chloridev 4at atmospheric pressure while another of such tanks is on stream.

The amount-of aluminum chloride that is dissolved in the solution tank is primarily dependent upon the temperature and the nature and amount of the butane stream. By simply controlling flow rates (pump 39) and temperature (heat exchanger 4|) we may closely controlthe amount of make-up aluminum chloride lwhich is thus charged to the system. 'I'he amount of this stream will usually be less than 50% of the total charging stock and for example may constitute about 10 to 30% of said charging stock. The temperature will usually be within the range oi about 100 to 400 lj'. or more, fory example, about 200 to 300 Il'.

The hydrogen chloride-rich charging stock from the base of absorber 21 is Dumped by Dump 44 through heater 4I and lines 40 and 41 to the base of the flrst reactor 48 at a pressure within the approximate range of 500 to 1500 pounds per square inch. for example. about 850 or 900 pounds per square inch. Hydrogen from source 40 (or from other sources that will hereinafter be described) is introduced by compressor 50 and line Il into line .48 in amounts within an approximate range of 100 to 300, for example about 200 cubici'eet per barrel of stock charged to the reactor (the hydrogen being measured at 60 F. and atmospheric pressure). The aluminum chloride solution from aluminum chloride solution tank 42 is introduced through line 52 (and pump 5I if further pressuring is required) into the reactor and the amount of aluminum chloride in solution in the butane may be about .l or 1 pound per gallon. Based on the total charging stock introduced into the reactor, the amount of within the approximate range of 1.2 to 1.7, and

which can be maintained during the reaction within the general vicinity of 1.5 by methods hereinafter described. The density of the light hydrocarbon charging stock is less than halt that 'of desired isomerization products. The complex carried into the soaking zone is as huid. as freshly prepared complex, i. e., it may be quite diilerent in its properties than the average complex in the initial reactor.

Drum 54 also serves the important function oi' removing any undissolved complex. separated complex being withdrawn from the base of the drum through line` I1 and returned without neo'essity of pumping. Complex removed "at this point helps to prevent fouling of the heat exchanger when clear products are `withdrawn .through line It, cooler (heat exchanger) 58, and

, the approximate range of 100 to 300, for example o the complex. 'I'he charging stock is chiey in the liquid phase but some of it will be vaporized by the upilowing gases which constitute about v stock per hour per volume of catalyst complex in the reactor. Higher space velocity may, of course,

be employed in this reactor, particularly where a plurality of reactors are employed in series.

Complex settles from the upowing reaction products in the top of the tower, and if desired the tower top can be enlarged' to provide increased settling area. We prefer, however, to withdraw the reaction products in the tower top through lines 54 and 55 to a soaking drum or warm settling chamber 56 which can be a horizontal or slightly inclined drum about 3 feet in diameter by about 10 feet long. We have discovered that there is a large amount of dissolved catalyst in the product at this point and this dissolved catalyst in the soaking drum may have a beneilcial eiIect on distribution of products. In other words, the isomerization equilibrium in drum 56 is not the same as in reactor 48 and this supplemental contact with dissolved catalyst in drum 58 may contribute to additional formation about 250 pounds per square inch. The cool settler may be a horizontal or slightly inclined drum about 5 or 6 feet in diameter and about 16 feet in length. Released gases leave the top of the cool settlerl through line 82 which discharges into line 2l. The reduced pressure and cooling eil'ects a considerable precipitation of catalyst material in the cooled settler and the precipitated catalyst material is withdrawn from the base of this settler by means of pump 63 in line B4.

The clear product, which is now substantially free from catalyst and which contains a decreased amount/ of hydrogen chloride, is withdrawn through line 85 andintroduced by pump 66 into hydrogen chloride stripping tower 61 which can be a column about 3 feet in diameter and about 33 feet in height. This stripping column can be provided with heating means 58 at its base and it can be operated at a pressure of about. 200 pounds per square inch with a top temperature, within the approximate range of F. to 150 F. and a bottom temperature within the approximate range of 300 F. to 400 F, The removed hydrogen chloride together with the released gases. such as hydrogen, methane, etc., is taken overhead through line 0I to line 3l.

The liquid from the base of the stripper is introduced at a low point in scrubbing tower 10 either directly through line 1I or through a cooler 12. scrubber 10 may be a tower about 4 feet in diameter, about 32 feet in height, and it can be provided with suitable baiiies, trays or bubble plates for effecting intimate contact of the upowing product with a concentrated caustic solution introduced through line 13. The upilowing neutralized products are washed free from caustic in the upper part of the tower by water introduced through line 14. Spent caustic solution is withdrawn from the base of the scrubber through line 15. The wash water can be withdrawn from a trapout plate .(not shown) above the point of caustic inlet, if desired.

The water-washed product can be' withdrawn as suchfor it can pass from the top of scrubber 10 through line 18 and heat'exchanger 11 to ytion tank 42 to supplement or supplant the bu- I tane from line 2| for making up the catalystA solution.

If desiredI a single isomate" fraction can be withdrawn from the base of the stabilizer through heat exchanger 11 and line 09. We can. however, withdraw only the heaviest isomate at this point and We can withdraw a light isomate as a separate fraction diagrammatically illustrated as a side stream through line 90. The isomate can be fractionated to insure the removal of any heptanes or heavier products which may -be formed and to obtain a product of desired Reid vapor pressure for blending in desired amounts with isooctane to make a super aviation fuel. A representative analysis of isomate producedin this system may be approximately as follows:

If neohexane' is a desired end product, it can be separately fractionated. and the other isohexanes can be recycled for the production of a further amount of neohexane.

Returning to the reaction system, we canemploy 'a second reactor 9| of similar design to the first reactor 43. Products from the first reactor instead of going to the warm settler through line 55 can pass through line 92 and heat exchanger 93 into the base of this second reactor 9|. The operating conditions in the second reactor can be substantially the same as in the first reactor, although we prefer to operate the second reactor at a lower temperature than the first reactor. Thus with the rst reactor at 300 F. or more, the second reactor can be at about 250 F. Products from the top of the second reactor pass through line 94 to soaking drum or warm settler 56 as hereinbefore described.

Instead of operating the 'reactors in series they may be operated in parallel by passing only a part of the charging stock through line 41 to the first reactor and bypassing the remainder of the charging stock through line 95 to the base of the second reactor. By means of this arrangement one reactor can be on stream while another reactor is standing by for'repair orreplacement of catalyst complex.

A In general the complex becomes more viscous with age and up to a certain point the catalyst becomes more active with increasing viscosity. These characteristics `are apparently determined to a certain extent by the hydrocarbon content of the complex, Fresh complex may contain up to about'3'1% hydrocarbon but ywhen complex' is made in situ in the presence of a large amount of aluminum chloride, and after the catalyst has been used for a period of time it may contain only 10 or 15% hydrocarbons, the balance consisting chiefly of aluminum chloride. We have found the catalytic activity of the complex can be'maintained substantially constant by withdrawing a portion of the catalyst from the base of the reactor at about the-same rate as an additional amount of aluminum chloride is added thereto, either in the form oit-relatively fresh complex or preferably in the form of an aluminum chloride-solution. Thus catalyst in the sec- Per cent by volume Isobutane 2 Isopentane 31 n-Pentane Cyclopenta-ney 3 2methylpentane 18 3-methylpentane 8 2,2-dimethylbutane 2 2,3-dimethylbutane v2 n-Hexane and heavier T 11 ond reactor can be withdrawn through line by means of pump 91 and either introduced through linee 90 and 99 to the first reactor, v

second reactor operates at a lower temperature than the first reactor when in series therewith, we prefer to introduce catalyst from the second reactor through lines 36, 99 and 09 to the'first reactor, and to remove catalyst through line |04 in the first reaction to the hydrogen chloride recovery drum |03 orto other recovery systems to be described later. For such operation Iwe prefer to introduce a part of all of the makeup aluminum chloride solution' through line 52a to the second reactor.

Settled catalyst complex from the soaking drum or warm settler 56 can be passed from line 51 through lines ||0 and to solution tank 42,.or through line |l2 to the second reactor 9|, or through lines ||4 and 99 to the rst reactor 48. Separated catalyst from cool settler 6| can be passed'from line 64 through lines |I3 and to Solution tank 42, or through line I2 to the second reactor 9| or through lines ||4 and 99 to the first reactor 48.

Spent sludge can be discarded from the system but we may either recover a part of the available aluminum chloride therein or introduce it into drum |03 and add to the sludge in this drum through line 5 a sufcient amount of water or sulfuric acid to effect recovery of anhydrous hydrogen chloride. Preferably,V the uncombined aluminum chloride is recovered by the method hereinafter to be described, and the remainder of the sludge is then introduced into drum |03 for the production of anhydrous hydrogen chloride. The recovered hydrogen chloride from drum |03 is passed through line ||B to line 3| and absorber 21.` The sulfuric acid sludge or cokey residue is withdrawn from drum |03 through line ||1. If wateris employed, it should be used in less than stoichiometric amounts in order that the recovered hydrogen chloride may be substantially anhydrols. The sludge will thereupon be converted into a cokey mass that can be removed from the drum by hydraulic or other conventionalv decoking means. A larger amount of Vanhydrous hydrogen chloride can be recovered by the use of sulfuric acid and the resulting sulfuric' acidi sludge can be charged to a conventional sludgecoking system for the recovery of sulfuric acid.

Instead of employing relatively pure hydrogen from source 49 we can obtain hydrogen from refinery gases which are rich in hydrogen. Such u gases can be introduced through line |20 to abdrogen will pass overhead through line to be picked up by compressor 50. The rich absorber oit will pass throughl line |20 and pressure reducing valve |21 to receiver |20 from whichthe hydrocarbon gases can be vented to fuel lines or other parts of the refinery through line |29 and the denuded oil can be returned by line |30, pump |23 and line |24 back to the top of the absorber.

As was previously mentioned, a part or all of the uncombined aluminum chloride can be recovered either from the spent catalyst withdrawn through lines |00 and/or |01 and from the catalyst recovered from warm settlerll and/or cool settler 6|. It appears that there il present in the spent catalyst from the reaction varying amounts of free aluminum chloride either in solution in the sludge or-held there by occlusion or other physical means, while much of the catalyst precipitated in the settlers is in the form of complex and/or aluminum chloride. The aluminum chloride can be recovered by dissolving it in a quantity of fresh hydrocarbon and after the soluble portion of the catalyst has been removed, the remaining sludge can be withdrawn for the production of hydrogen Chlo--V ride, as previously described, or for other purposes. actors 48 and/or 9| can be directed via valved lines |31 and |38 respectively to line |33 leading to catalyst recovery tank |40. A portion of the butane from line 2| is introduced into recovery tank |40 via valved line I4| and line |42. Means (not shown) are provided for the intimate mixing of the sludge and the butane, preferably at elevated temperatures, such as those set forth for solution tank 42. The butane A containing dissolved aluminum chloride is Accordingly, the spent sludge from re' 4 the stream passes through lines 2| and 40 to solution tank 42. When using a stock of this type butane from line 30 is not necessarily reline or directed to the hydrogen chloride recovery system by opening valve |46 in line |41 which joins line |02.

'I'he catalyst which has settled out in warm settler 56 and/or cool settler 6| can' also be directed to a recovery system or returned directly to the reactors, usually in the form of a complex or a slurry in a portion of the product from the settlers. At the temperatures previously disclosed the precipitated -catalyst is redissolved in the butane and can be thereupon directed touse in the reaction zones. The recovery tank for the re-solution of aluminum chloride from the settlers can be' the same -as that used for the recovery of free aluminum chloride from the sludges as shown, or a separate recovery tank (not shown) can be used.

A further alternative procedure is to withdraw butane and light naphtha as a single stream and to direct a portion of this stream through absorber 21 wherein it is saturated with hydrogen'chloride and another portion of the stream through heat exchanger 4|t0 catalyst solution tank 42 wherein a sumcient amount of aluminum chloride is dissolved to supply the necessary make-up aluminum chloride to the system. In this system light naphtha will not be withdrawn as a side stream through line 25 but be included in the overhead from fractionator i3 and recovered from reflux drum I8. The major part of the combined butane and light naphtha stream is directed via pump |0, line |48 pended claims.

cycled.

Y Another alternative is toy use only the light naphtha stream for dissolving the make-up catalyst, eliminating substantially all butane from the system. When employing this process, the butane from reflux drum Il will be withdrawn through line 22, and only the pentane-hexane stream from line V25 employed. 'I'he light naphtha is preferably used at elevated temperatures for making up'the solution and temperatures within the range from about F. to 400 F., preferably about 250 F.. are suitable.

It may be desirable to supply additional aluminum chloride in the form of a hydrocarbon complex or as a slurry from an outside source (not shown). 'I'he aluminum chloride solution can be prepared as previously described from the fresh naphtha feed and introduced in a. separate stream through line 02.

The introduction of make-up catalyst in the form of a solution is a distinct advance over the prior art. By this means it is possible to obtain an even and regulated flow of make-up aluminum chloride to the reaction system so that the comvplex is formed in situ and can be adjusted according to the amount of catalyst necessarily withdrawn. Moreover, the introduction is smooth and easy with no plugging of lines and no settling out of catalyst as is often the case when a slurry type addition is employed. More intimate mixing is possible by the use of a solution of aluminum chloride than is possible by the use of a slurry and even by the use of a preformed complex since there is is no stratification within the reactor or in the lines and turbulence is only necessary in order to insure intimate contact between the reactants and the catalyst.

There is an added advantage to the use of butane as a solvent for the introduction of catalyst apart from the fact that much larger quantitles of aluminum chloride can be dissolved and that is the fact that higher temperatures, particularly temperatures upwards of 350 F. can be safely used without excess degradation of the charging stock to light gases. The butane apparently acts as a buffer gas and prevents or minimizes the production of additional quantities of gaseous hydrocarbons such as are generally formed at the higher temperatures above recited. This advantage is even more pronounced if a feed stock containing heptanes and higher paralns is used.

Although we have directed our description to a process involving the isomerization of a light naphtha fraction it is possible to use this system for the isomerization of normal butane to isobutane with no light naphtha present. Suiiicient make-up catalyst may be dissolved in the incoming butane feed to supply the needs of the reaction. Obviously, variations in temperature and pressure from those described lwould be requireclv and it is usually desirable to carry out butane isomerization in the absence of hydrogen or at least under very low hydrogen pressures,

Although we have described our process with reference to one preferred embodiment thereof, it should be realized that this is by way of illustration and not by way of limitation and that' our invention is limited only as set forth in the ap- We have also omitted various details the necessity for which would be obvious to one skilled in the art wishing to practice our invention.

We claim:

1. In a-process for isomerizing paraffin hydrocarbons of the butane to hexane boiling range by means of a liquid aluminum chloride-hydrocarbon complex catalyst wherein the activity of the catalyst in a reaction zone is maintained by the additionv of make-up aluminum chloride thereto, the method of operation which comprises maintaining a mass of aluminum chloride-hydrocarbon complex in said reaction zone up to a high level which is spaced from the topof said zone, heating a portion of a charging stock consisting chieily of paraffin hydrocarbons of the butane to hexane boiling range in the absence of hydrogen chloride to a temperature within the approximate range of 200 to 300 F. and passing the heated portion through a mass of solid aluminum chloride for dissolving a regulated amount of aluminum chloride, absorbing hydrogen chloride in another portionof said charging stock and heating said other portion to an isomerization temperature below 300 F., introducing both portions of said charging stock into said mass of 'complex catalyst and passing said charging stock upwardly throughsaid mass under isomerization conditions at a temperature below 300 F., separating hydrocarbons from complex in the upper part of said zone,`removing hydrocarbons from said zone at a point above the level of the complex mass of catalyst therein, separating catalyst material and hydrogen chloride from the removed hydrocarbons, recycling separated hydrogen chloride without contacting it with said solid aluminum chloride, and introducing an amount of aluminum chloride by said solution step to provide an over-all make-up of more than .1 but less than 4 pounds of aluminum chloride per barrel of stock charged.

2. An isomerization process which comprises preparing a paraiilnic hydrocarbon charging stock of the butane to hexane boiling range which contains substantially no oleflns, less than ing one portion of said charging stock in the absence oi hydrogen chloride to a temperature within the approximate range of 200 F. to 300 F. and passing s'aid heated portion through a mass of solid aluminum chloride for dissolving a controlled amount of 'aluminum chloride, ab sorbing in another portion of the charging stock an amount of hydrogen chloride within the approximate range of 2 to 10% by weight based on total stock charged and eecting said absorption in said other portion of the charging stock before said other portion is heated, subsequently heating said other portion of the charging stock containing absorbed hydrogen chloride to an isomerization temperature but below 300 F., in-

troducing both portions into a deep body of liquid aluminum chloride-hydrocarbon catalyst complex in a conversionzone, passing said charging stock upwardly through said body of complex under isomerization conditions at a temperature below 300 F., separatingv hydrocarbons from complex in the upper part of said zone, removing hydrocarbons from the upper part of said zone at a point above the top of thebody o i complex therein, separating catalyst material and hydrogen chloride from the removed hydrocarbons, recycling separated hydrogen chloride to said absorbing step, and introducing an amount of -aluminum chloride by said solution step to provide an over-all make-up of more than .-1 but lessv than 4 pounds of aluminum chloride per barrel of stock charged.

v3. The method of claim 1 wherein the portion of the charging stock contacted with solid aluminum chloride consists vessentially of nbutane. Y 4. The method of claim 2 wherein the portion of the charging stock in'which hydrogen chloride is absorbed consists essentially of hydrocarbons of the pentane-hexane boiling range.

EDMoND L. noUvrLLE. BERNARD 1..-EvERmG. 

